Method and recovery of benzene and toluene

ABSTRACT

A METHOD OF RECOVERING BENZENE AND TOLUENE FROM NORMALLY LIQUID HYDROCARBON MIXTURES CONTAINING IN EXCESS OF 50% OF BENZENE AND TOLUENE, TOGETHER WITH ACYCLIC MONO-OLEFINS, ACYCLIC DIOLEFINS, CYCLO-PARAFFINS, CYCLIC MONO-OLEFINS, CYCLIC DIOLEFINS, AND IN SOME CASES, SULFUR AND NITROGEN ORR OXYGEN COMPOUNDS, INCLUDING DISTILLING THE MIXTURE OF RECOVER A FRACTION BOILING IN THE BENZENE-TOLUENE RANGE, SUBJECTING THIS FRACTION TO A HYDROGEN TREATMENT IN THE PRESENCE OF A CATALYST AND UNDER CONDITIONS TO DEHYDROGENATE ACYCLIC DIOLEFINS AND HYDROGENATE ACYCLIC MONO-OLEFINS, POLYMERIZING THE HYDROGENATION PRODUCTS, IF DESIRED, DISTILLING THE POLYMER PRODUCT TO RECOVER A FRACTION BOILING IN THE BENZENE-TOLUENE RANGE, SUBJECTING THE BENZENE-TOLUENE FRACTION TO A HYDROGEN TREATMENT IN THE PRESENCE OF ACATALYST AND UNDER CONDITIONS SUFFICIENT TO CONVERT NONAROMATIC CYCLIC COMPOUNDS TO AROMATIC COMPOUNDS AND CRACK OTHER NONAROMATIC COMPOUNDS OF LOWER BOILING POINT.

April 23, 1974 s. M. Ko'vAcH ET AL Original Filed Sept. 28, 1966 METHOD AND RECOVERY OF BENZENE AND TOLUENE Sheets-Sheet l Feed Mlxure BPM/@M114 ATTORNEY ET AL 3,806,553

2 Sheets-Sheet 2 JNVENTORS S, M. KOVACH METHOD AND RECOVERY OF BENZENE AND TOLUENE April 23, 1974 Original Filed Sept. 28, 1966 Uni'ted States Patent 3,806,553. Patented Apr. 23, 1974 U.S. Cl. 260-674 R 4 Claims ABSTRACT F THE DISCLOSURE A method of recovering benzene and toluene from normally liquid hydrocarbon mixtures containing in excess of 50% of benzene and toluene, together with acyclic mono-olens, acyclic diolefins, cyclo-paraflins, cyclic mono-olens, cyclic diolefins, and in some cases, sulfur and nitrogen or oxygen compounds, including distilling the mixture to recover a fraction boiling in the benzene-toluene range, subjecting this fraction to a hydrogen treatment in the presence of a catalyst and under conditions to dehydrogenate acyclic diolens and hydrogenate acyclic mono-olens, polymerizing the hydrogenation product, if desired, distilling the polymer product to recover a fraction boiling in the benzene-toluene range, subjecting the benzene-toluene fraction to a hydrogen treatment in the presence of acatalyst and under conditions suicient to convert nonaromatic cyclic compounds to aromatic cornpounds and crack other nonaromatic compounds of lower boiling point.

RELATED APPLICATIONS This application is a division of our copending application, Ser. No. 582,652, tiled Sept. 28, 1966, and now abandoned.

The present invention relates to the separation of aromatic hydrocarbons. In a more specific aspect, the present invention relates to the separation of normally liquid hydrocarbon mixtures containing at least about 50% by volume of aromatics. In a still more specific aspect, the present invention relates to the recovery of benzene and toluene from normally liquid hydrocarbon mixtures containing at least about 50% by volume of aromatics.

Normally liquid hydrocarbon mixtures containing in excess of about 50% of benzene and toluene are commercially available from a variety of sources. These mixtures contain, in addition to the benzene-toluene, acyclic mono-oletins, including, secondary and tertiary mono-olens of normal and branch chain structure, acyclic dioletins, cyclo-parains, cyclic mono-olehs and cyclic diolefins. Many of these hydrocarbon mixtures will also contain varying percentages of sulfur and some of them will also contain nitrogen or oxygen compounds.

Processes for the production of aromatic-rich hydrocarbon mixtures include the pyrolysis or cracking of crude petroleum, or fractions thereof, containing at least two carbon atoms such as ethane, propane, propylene, butane, natural gasoline, light straight run gasoline, straight run naphtha, kerosene, light cycle oils produced in the cracking of gas oils to produce gasoline, straight run gas oil, etc. These processes are carried out with or without the aid of catalysts and in the presence or absence of steam. Generally, the reaction temperature is in the neighborhood of about 1350 to 1550 F. and the pressure may be from 0 to 50 p.s.i.g. or higher. Such pyrolysis operations are generally carried out primarily for the production of ethylene. The pyrolysis of hydrocarbons to produce ethylene results in a normally gaseous product containing unsaturated hydrocarbons, including ethylene; normally liquid hydrocarbons rich in unsaturated hydrocarbons, including, oletns and diolens of varying boiling points and structures and various aromatic hydrocarbons; as well as tar. The normally liquid hydrocarbons and tar obtained from this process are considered byproducts. These byproducts are removed by rapidly cooling the pyrolysis products, usually by quenching with water, to a temperature of about 400 F. A viscous tarry material condenses out of the gas during the quenching operation. Gases from the quenching operation are then compressed and cooled and a liquid material boiling between about and 360 F. condenses out of the gases during the compression-cooling operation. This last material is known as dripolene.

Dripolene fractions obtained in the manner set forth normally contain about two-thirds aromatics and onethird non-aromatics, the latter being mostly oleins and diolens. In a typical situation dripolene will contain about 64% aromatics consisting essentially of benzene and toluene and about 36% non-aromatics including acyclic mono-olens, including, secondary and tertiary monoolefins of normal and branch chain type, acyclic diolens, cyclo-paraflins, cyclic mono-olelins, cyclic diolefins, and, in some cases, oxygen compounds and sulfur compounds.

It has heretofore been suggested that the removal of aromatics from dripolene and other aromatic-containing mixtures can be accomplished by selective adsorption on solid materials, such as silica-gel, activated alumina, etc. However, such selective adsorption of aromatics is relatively ineifective since diolefins, present in the mixture, behave quit similarly to the aromatics. The other approach is, of course, the removal of the unsaturated compounds from the aromatics. Such removal of unsaturates has heretofore been practiced by a number of processing schemes, for example, hydration, halogenation, hydrohalogenation, carbonization, hydrogenation followed by glycol-water extraction, ozonolysis, polymerization, etc. However, because of the high concentrations of unsaturates in mixtures of aromatics, such as dripolene, processing costs and aromatic losses due to reactivity of the aromatic compounds has generally confined the prior art treatments, at least commercially, to hydrogenation and polymerization.

Under the conditions normally employed in the hydrogenation of aromatic-rich mixtures, such as dripolene, including, using a palladium catalyst for diolen conversion followed by a second hydrogenation and finally the glycol-water extraction of the product, the cost of hydrogenation is usually quite high, since large quantities of hydrogen are necessary, and the glycol-water extraction alone is an expensive proposition. In addition, high pressures and high hydrogen to hydrocarbon ratios must be employed to produce an essentialy olefin-free product. Under these conditions, diolens normally undergo a Diels-Alder reaction to produce polymers, gums and resins. Further, when a palladium catalyst is used it is readily poisoned by feed mixtures having normal sulfur contents and hydrogen containing small amounts of free hydrogen sulfide.

fWhere polymerization is utilized for the separation of unsaturates from aromatic mixtures, typical polymerization conditions lead to undesirable losses of aromatics. This is caused by the alkylation of the aromatics, which takes place with approximately the same ease as the polymerization.

Another convenient source of aromatic hydrocarbons, particularly benzene and toluene is an aromatic light oil obtained by the high temperature carbonization of coal. The products of this carbonization are coke and the aromatic light oil. This light oil contains relatively high concentrations of unsaturates, sulfur and nitrogen compounds in addition to benzene, toluene and xylene. For example, a coal tar light oil boiling in the range of about 150-300 F. will contain, roughly, 86% benzene-toluene and about 14% unsaturates and impurities. The non-aro- -matics include acyclic mono-olefins, including, secondary and tertiary mono-oleiins of normal and branch chain ty-pe, acyclic diolens, cyclic paraflins, cyclic mono-olefins, cyclic diolens and substantial amounts of sulfur and sul-fur compounds as well as nitrogen compounds.

Heretofore, the recovery of aromatics from coal tar light oils, particularly the recovery of benzene, has been effected by acid washing with sulfuric acid, benzene crystallization followed by acid washing, hydrotreating followed by solvent extraction With an aromatic-selective solvent or hydrotreating followed by high severity hydrocracking. Each of these treatments has been found to have its own peculiar drawbacks. For example, acid treating alone results in a benzene product of high purity but having a high thiophene content. Benzene crystallization followed by acid washing will produce a product of high purity having a low thiophene content; but the recoveries of benzene are low. Hydrotreating followed by solvent extraction produces a high purity benzene product but is a costly process and coking problems in the preheater to the hydrotreater exist. The nal technique, of hydrotreating followed by high severity hydrocracking, is a costly proposition because of the high pressures and high temperatures involved and at times such treatment does not produce specification grade benzene.

In accordance with the present invention it has been surprisingly discovered that olens may be readily removed frorn liquid hydrocarbon mixtures containing, in excess of about 50% aromatics, by selectively distilling the mixture to recover a heart-cut rich in desired aromatics, selectively treating the heart-cut fraction to convert the acyclic diolefins and at least a portion of the acyclic mono-olens, and, finally, selectively treating at least a portion of the product of the rst selective treatment to convert the cyclic compounds to aromatics and any remaining nonaromatic compounds to compounds having boiling points differing from the boiling range of the desired aromatics. v

The details of the present invention can best be understood and exemplified by reference to the following description when read in conjunction with the drawings, wherein:

FIG. 1 is a flow diagram of one form of the present invention; and

FIG. 2 is a flow diagram of another form of the present invention.

It has been discovered quite surprisingly, in Vaccordance with the present invention, that normally liquid hydrocarbon mixtures, containing above about 50% of aromatics, can be treated to remove the aromatics therefrom by selective distillation to remove a heart-cut fraction containing the desired aromatics, thereafter selectively hydrogenating the heart-cut fraction to dehydrogenate cyclic dioleiins and saturate acyclic olens, and selectively treating the hydrogenation product to dehydrogenate cycloparans and cyclic mono-olefns and crack saturated paraiiins to lower boiling parafiins boiling below the boiling range of the desired aromatics.

According to one specific technique for separating benzene and toluene from highly aromatic feed mixtures, a select heart-cut of the feed mixture is subjected to hydrogenation under highly select conditions and, thereafter, the product is subjected to a hydrocracking or reforming operation.

The novel features of this technique will be best illustrated by reference to FIG. 1 of the drawings.

In accordance with the drawing, a highly-aromatic feed mixture is introduced through line from a source not shown. The feed mixture from line 10 passes to a distillation column 12 in which a heart-cut material, such as, a benzene-toluene concentrate, is separated and discharged through line 14. A benzene-toluene concentrate has a boiling range of about to 250 F. and preferably to 245 F. Feed material boiling below about 140 F., and preferably below 160 F., is discharged through line 16 while feed mixture boiling above 250 F., and preferably above 245 F., is discharged through line 18. The high and low boiling portions of the feed mixture can be further treated to produce novel and valuable products. However, the techniques for producing these products and the products of these techniques form no part of the present invention. It has been found that feed mixtures of highly aromatic character, such as, dripolene, often contain oxygenated compounds. Specifically, dripolene has been found to contain about 2.5% by weight of esters such as ethyl acetate. It has been found highly desirable, in the present invention, to remove these oxygenated compounds before subjecting the heartcut fraction to hydrogenation treatment. Accordingly, the concentrate from line 14 may be fed through valve 20 and line 22 to an oxygen removal unit 24. The oxygen removal unit 24 will selectively remove oxygenated compounds or substantially reduce their volume. It has been found that the high volumes of oxygen compounds, previously mentioned, can be reduced to about 0.3% by weight by such a treatment. Selective adsorption utilizing a 10x- 13x molecular sieve is preferred. It should also be recognized, however, that other appropriate means for removing oxygenated compounds may be employed such as water washing or processing over selective adsorbents such as silica gel. To the extent an oxygen removal unit 24 is utilized, valve 25 in line 14 will be closed when valve 20 to unit 24 is open. The treated concentrate from unit 24 is discharged through line 26. From line 26 the concentrate passes to hydrogenation unit 28. Hydrogenation unit 28 is operated under highly select conditions to hydrogenate acyclic olens and dehydrogenate cyclohexadienes to benzene and toluene. More specifically, the hydrogenation may be carried out at temperatures between about 250 and 300 F., a pressure of about 0 p.s.i.g. and a liquid hourly space velocity between about 0.1 to 10, and, preferably, between about 0.5 to 2. Hydrogen, at a rate of about 100 to 1000 cubic feet per barrel of feed, and, preferably, between about 0.5 to 2. Hydrogen, vat a supplied to hydrogenation unit 28 from an external source not shown, through line 30. Suitable catalysts include 0.1 to 10% platinum or palladium on alumina, spent reforming catalysts, etc. The following table gives specific examples of the treatment of benzene-toluene concentrates which have not been pretreated to remove oxygenated compounds and one Which has been passed through a molecular sieve. The specific feed in this particular instance was a heart-cut, obtained by distilling raw dripolene, originally containing 52.4% benzene and 11.6% toluene, to obtain a fraction boiling between 160 and 245 F. This fraction represented 68.2% by volume and included 9.2% of compounds boiling below benzene, 76.1% benzene, 14.1% toluene and 0.5% of compounds boiling above toluene and 69.5% by volume and included 14.5% of compounds boiling below benzene, 68.5% benzene, 16.1% toluene and 1.0% of compounds boiling above toluene, in runs 1 and 2, respectively.

It is apparent from the data presented above that the presence of oxygenated compounds in the feed to hydrogenator 28 affects the dehydrogenation activity of the catalyst. In the presence of esters or other oxygenated compounds, the cyclohexadienes are apparently hydrogenated thus increasing hydrogen consumption, while in the absence of oxygenated compounds cyclohexadienes are dehydrogenated to aromatics thus producing hydrogen which in turn hydrogenates other unsaturates. Thus, where oxygenated compounds are present it is highly desirable that the feed be pretreated to remove these compounds prior to hydrogenation.

The hydrogenated product is discharged through line 32. This hydrogenated product may be treated by a variety of hydrocracking and/ or reforming operations under select conditions. Since the hydrogenated product is to be subjected to hydrocracking or reforming to convert nonaromatics to lower boiling parains and the cyclohexanes to benzene and toluene, the product need not have a bromine number which is nil. In addition, under the conditions herein set forth, a net increase in aromatics results, as evidenced by the fact that the aromatic recovery is greater than that originally present in the feed.

In one variation of this techninue the hydrogenated product is passed through line 34 and valve 36 to hydrocracking unit 38. Hydrocracking unit 38 is operated under low severity hydrocracking conditions. More specifically, hydrocracking is carried out at a temperature between about 800 and 1200 F., and, preferably 950 to 1100 F., at a pressure between about 0 to 100 p.s.i.g., and, preferably, 50 to 200 p.s.i.g., and at a liquid hourly space velocity (LHSV) of about 0.1 to 10, and, preferably, 0.5 to 2. Hydrogen, at a rate of about 100 to 10,000 cubic feet per barrel of feed, and, preferably, 1000 to 2000 cubic feet per barrel, is supplied from an external source, not shown, through line 40. The hydrocracking catalyst may be any active hydrogenation-dehydrogenation metal on an acidic solid oxide support. Preferably, the catalyst is 7% nickel on silicaalumina. Hydrocracked product is discharged through line 42 from whence it passes to distillation column 44 for further processing to recover the desired products. In distillation column 44 the material is separated to remove material boiling below about 176 F., which is discharged through line 46. Toluene is separated in column 44 and discharged through line 48. The intermediate cut of benzene is discharged through line 50. Preferably the separated benzene is passed to clay treating unit 52 where it is treated by a conventional hot clay treatment to remove trace impurities. The clay treated product is then discharged through line 54.

The following table shows the results of hydrocracking the hydrogenation product of Run #l of Table I.

TABLE II Catalyst 7 Ni/SiOz-Al203 Temperature, F. 1050 Pressure, p.s.i.g 100 LHSV 0.75 H2,ft.3/bbl 1400 Product (160-245 F.) percent by volume:

Recovery 95 Benzene 1.2

Benzene 80.4

Toluene 16.5

Toluene 1.9

In the above run it is to be observed that conversion of nonaromatics was not complete. This can be rectified by utilizing a slightly higher temperature or by employing a second hydrocracker at the same temperature set forth in the example. Specifically, valve 82 in line 42 may be closed to pass the product of hydrocracker 38 through line 84 and valve 86 to line 32. From line 32 the feed passes through line 88 and valve 90 to second hydrocracker 92. Hydrocracked product from unit 92 is discharged through line 94 to line 42. Hydrogen is supplied to hydrocracker 92, from a source not shown, through line 96.

The hydrogenated product from line 32 may be passed through line 56 and valve 58 to reforming unit 60. In reforming unit 60, the hydrogenated material is subjected to selective, high severity reforming conditions. Specifically, reforming is carried out at a temperature of about 800 to 1000 F., and, preferably, 900 to 950 F., a pressure of about 0 to 1000, and, preferably, 50 to 200 p.s.i.g. and at a liquid hourly space velocity (LHSV) of about 0.1 to 20, and, preferably, l to 10. Hydrogen, from an external source not shown, is supplied to reformer 60 through line 62. Hydrogen rates of to 10,000 cubic feet per barrel of feed and preferably 1000 to 4000 cubic feet per barrel are used. The reforming catalyst is preferably one containing 0.1 to 10% by weight of platinum or palladium on alumina. Reformed product is discharged through line 64 and may then be passed to distillation column 44 through line 42.

The following table shows the results of treating the hydrogenation product of Run #l of Table I, in accordance with this technique.

TABLE III Catalyst 0.6 Pt/Al203 Temperature, F. 940 Pressure, p.s.i.g 100 LHSV 4 H2, ft.3/bbl. 2800 Product (160-245 F.) percent by volume:

Recovery 94 Benzene 0.5

Benzene 81.4

Toluene 16.9

Toluene 1.3

In a specific technique, as illustrated by FIG. 1, the feed mixture of highly aromatic material is passed through line 10 to distillation column 12. In column 12 the broad boiling range material is separated to remove a heart cut or benzene-toluene concentrate, which is discharged through line 14. The benzene-toluene concentrate will have a boiling range of about to 250- F. and, preferably, to 245 F. Feed mixture below about 140 F., and, preferably, below 160 F., is discharged through line have a boiling range of about 140 to 250 F. and, preferably 245 F., is discharged through line 18. The higher and lower boiling cuts of the feed mixture may be further processed to produce valuable products. However, these processing techniques and the products thereof form no part of the present invention. The concentrate in line 14 is passed to hydrogenation unit 28 by opening valve 25. In hydrogenation unit 28, the aromatic concentrate is subjected to a hydrogenation-dehydrogenation reaction under select conditions and utilizing. a select catalyst system. More specifically, the concentrate is dehydrogenated to convert cyclohexadienes to aromatics and hydrogenated to convert the remaining diolefins to saturated compounds. The hydrogenation conditions include a temperature of about 75 to 400 F., and preferably, 100 to 300 F., a pressure of about 0 to 3000, and, preferably, 0 to 1000 p.s.i.g. and a liquid hourly space velocity of about 0.1 to 10, and, preferably, 0.5 to 2. Hydrogen is supplied to the hydrogenation unit, from an external source not shown, through line 30 at a rate of about 100 to 1000 cubic feet per barrel of feed. The novel catalyst employed in this instance is rhenium oxides or rhenium sulfides on alumina. This catalyst is insensitive to poisoning by materials such as oxygen or sulfui compounds. In fact, in the presence of sulfur compounds and hydrogen, rhenium heptoxides, pentoxides, etc. are converted to a suldes (heptasulde, pentasuliide, etc.) which are as active as the parent compound. Of course, the catalyst could be converted to the sulfide prior to use, by treatments, such as, contact with hydrogen sulfide, etc. The rhenium should be present in amounts of 0.1 to 10% by weight based on the total weight of catalyst. The results of distillation of raw dripolene to obtain a heart-cut yboiling between 160 and 245 F. followed by hydrogenation is shown in the table below. The heart-cut fraction represented 69.5% of the dripolene and included, 10.6% of compounds boiling below benzene, 74.8% benzene, 14.1% toluene and 0.5% of compounds boiling above toluene.

TABLE IV Catalyst 0.5% Rhenium/AIZO'S Temperature, F 300 Pressure, p.s.i.g LHSV 0.75 H2, ft. 3/bbl 200 Product (160-245 F.) percent by volume:

Recovery 100 Benzene 8.8

Benzene 74.2

Toluene 15.18

Toluene 1.2

While the above treatment resulted in the hydrogenation of acyclic diolens as well as dehydrogenation of cyclohexadienes, operation at a pressure about 100 p.s.i.g. higher, or between 10041000 p.s.i.g., will hydrogenate all acyclic olefins.

Hydrogenated aromatic concentrate is discharged through line 32 from whence it may be subjected to several treating techniques to convert nonaromatics to lower boiling paraffins and cyclohexanes to aromatics.

Specifically, one technique for treating hydrogenated aromatic concentrate is to pass the hydrogenated product through line 34 and valve 36 to hydrocracking unit 38. In hydrocracking unit 33 the feed is subjected to a low severity hydrocracking at a temperature of about 800 to 1200 F. and preferably 950 to 1l00 F., a pressure of O to 1000 p.s.i.g., and, preferably, 50 to 200 p.s.i.g., and a liquid hourly space velocity of about 0.1 to 10, and, preferably, 0.5 to 2. Any active hydrogenation-dehydrogenation metal on an acid oxide may be used as a catalyst. Hydrogen, at a rate of 100 to 10,000 cubic feet per barrel, and, preferably, 1000 to 2000 cubic feet per barrel, is supplied to hydrocracking unit 38 through line 40, from an external source not shown. Hydrocracked product is discharged through line 42 and passes to distillation column 44. In distillation column 44 material boiling below about 176 F. is separated and discharged through line 46. Toluene is separated and discharged through line 48 and benzene is discharged through line 50. The benzene may be further clarified to remove trace impurities, by a hot clay treatment of conventional character, in clay treating unit 52. The purified benzene is discharged through line 54.

It should be recognized that the rhenium on alumina catalyst may be used in both the hydrogenation and hydrocracking operations.

Instead of hydrocracking the hydrogenated product the product passing through line 32 may be passed through line 56 and valve 58 to reforming unit 60. In unit 60 the feed is subjected to reforming at a temperature of about 800 to 1000" F., and, preferably, 900 to 950 F., at a pressure of about 0 to 1000 p.s.i.g., and, preferably, 50 to 200 p.s.i.g., and at a liquid hourly space velocity (LHSV) of A0.1 to 20, and, preferably, 1 to 10. The reforming catalyst may be platinum or palladium, in an amount of 0.1 to by weight, on alumina. Hydrogen, from an external source, not shown, is supplied to reformer 60 through line 62 at a rate of about 100 to 10,000 cubic feet per barrel, and, preferably, 1000 to 4000 cubic feet per barrel. Reformate is discharged through line 64 where it passes to line 42 .and thence to distillation column 44. l

It should also be recognized that the rhenium on alumina catalyst can be utilized in both the hydrogenation and the reforming steps.

The hydrogenated product from line 32 may be passed through line 66 and valve 68 to polymerization unit 70. In polymerization unit 70, the hydrogenated product is subjected to a mild polymerization treatment. This polymerization treatment is carried out at a temperature of about 0 to 300 F., and preferably, 100 to 200 F., a pressure of about 0 to 1000 p.s.i.g., and, preferably, 50 to 200 p.s.i.g., and at a liquid hourly space velocity (LHSV) of about 0.1 to 10, and, preferably, 0.5 to 20. Any acidic, solid oxide catalyst may be utilized for the polymerization reaction. However, it is also to be noted that the previously-mentioned rhenium on alumina can be utilized in both the hydrogenation and polymerization treatments. Polymerization product from column 70 is discharged through line 72. This product is passed to distillation column 74 where a resinous material boiling above 250 F., and, preferably, above 245 F., is separated and discharged through line 76. The lower boiling portion of the polymerization product is discharged from column 74 through line 78 and valve 80. The product from line 78 is passed to line 32 from whence it passes to either hydrocracking unit 38 or reforming unit 60. From thi-s point the material is treated in substantially the same Way as that previously described for the direct treatment of the hydrogenated product by hydrocracking or reforming. The rhenium on alumina catalyst can also be used in hydrocracking or reforming of the distilled polymerization product.

Highly aromatic feed mixtures, which contain high concentrations of unsaturates, sulfur and nitrogen compounds, such as, coal tar light oil, can be treated by subjecting the aromatic-rich material to selective distillation, hydrogenation and two-stage, low-severity hydrocracking.

More specifically, crude coal tar light oil is fed to the system through line 10, from a source not shown. The feed mixture then passes to distillation column 12, where it is separated to remove a .benzene-toluene concentrate or heart-cut boiling between about and 250 F. and, preferably, between and 245 F. This material is discharged through line 14. Any lower boiling material below about 140 F., and, preferably, below 160 F. is discharged through line 16 and material boiling above about 250 F., and, preferably, 245 F. is discharged through line 18. The material in line 14 is passed through valve 26 to hydrogenation unit 28. In hydyrogenation unit 28, the concentrate is subjected to selective hydrogenation-dehydrogenation conditions. Specifically, the material is processed at a temperature of about 75 to 400 F., and, preferably, 100 to 300 F., at a pressure of about 0 to 3000 p.s.i.g., and, preferably, 0 to 1000 p.s.i.g., and at a liquid hourly spaced velocity (LHSV) of about 0.1 to 10, and, preferably 0.5 to 2. Hydrogen is supplied to the system, from an external source not shown, through line 30 at a rate of about 100 to 1000 cubic feet per barrel of feed. The catalyst, employed in the treatment of such high sulfur materials, is preferably a sulfur-insensitive catalyst and low temperatures are maintained, as previously indicated. Specifically, a catalyst of about 0.1 to 10% by weight of rhenium in its oxide or sulfide form on an inert support, such as alumina, should be utilized. In the presence of hydrogen and sulfur, rhenium heptoxide, pentoxide, etc. will be converted to sulfides (heptasulfides, penta-sulfides, etc.) without in any way affecting its catalytic activity. Consequently, this catalyst may be pretreated with materials such as hydrogen sulfide to produce the sulfide form of catalyst before it is utilized in the hydrogenation treatment. It has been found that when utilizing this catalyst under the mentioned conditions the treatment can be carried out with no noticeable coking in the hydrogenation unit or its preheater. The hydrogenated product is discharged through line 32 and then to line 34 through valve 36 to hydrocracking unit 38. The product from hydrocracking unit 38 is discharged through line 42. By closing valve 82 in line 42, the hydrocracked product may be fed through line 84 and valve 86 to line 32. From line 32 it may then be passed through line 88 and valve 90 to second hydrocracking unit 92. Hydrocracked product from second unit 92 is discharged through line 94 where it passes to line 42 and thence to distillation column 44. In distillation column 44 the product is separated to remove material boiling below about 176 F., which is discharged through line 46. Product toluene is removed through line 48 and product benzene through line 50. The benzene, preferably, is hot claytreated in clay treating unit 52.' Puried benzene is discharged through line 54.

Both hydrocracking units 38 and 92 effect low severity hydrocracking of the material fed to them. Specifically, hydrocracking is carried out at temperatures of about 800 to 1200 F., and, preferably, 950 to 11005 F., at a pressure of about to 1000 p.s.i.g., and, preferably, 50 to 200 p.s.i.g., and at a liquid hourly space velocity (LHSV) of about 0.1 to 10, and, preferably, `0.5 to 2. Hydrogen is supplied to hydrocrackers 38 and 92, from external sources not shown, through lines 40 and 96, respectively. The hydrogen to hydrocarbon mole ratio 1-20/ 1, and, preferably, 1-5/1. The catalyst in the hydrocracking units may be any active hydrogenation metal such as nickel, platinum, cobalt, molybdenum, tungsten, etc. on an acidic solid oxide support, such as, silica-alumina, boria-alumina, silica-magnesia, fluoride promoted alumina, etc. The following table gives the results obtained when a coal tar light oil was processed as set forth above. The crude coal tar light oil was first distilled to obtain a fraction boiling between 150 and 240 F. which represented 88% of the crude starting material and contained, 2.0% of compounds boiling below benzene, 79.8% benzene, 18.4% toluene, 0.2% of compounds boiling above toluene, 5000 parts/million (ppm.) sulfur and 400 p.p.m. N.

TABLE V Hydro- Hydro- Hydrogenation cracking cracking 0.5 Re/ NiS-WSz/ NiS-WSQ/ Catalyst A1203 Slot-A1203 Sich-A1203 Temperature, F--. 250 950 950 Pressure, p.s.i. 100 100 100 LHSV 0.8 0.8 0.8 H2 mole/mole. 1/4 1/1 1/1 Process time, hrs 16 12 12 Prmlct (liquid): t by ecovery percen VOlllIIlo 100 95 Benzene, percent by volume 2. 3 0. 9 Benzene, percent by volnme. ..l 78.8 78.7 Toluene, percent by volume 18. 2 19. 2 Toluene, percent y volume 0.7 1. 2 S, p.p.m 3,500 100 N, p.p m 10 Instead of using two hydrocracking units in the technique described above, one of the hydrocrackers may be replaced by an appropriate hydrotreater. The hydrotreater preferably replaces hydrocracker 38, or the first of the two hydrocrackers. Thhe hydrotreater has the advantage, over the hydrocracker, that it is a more eicient sulfur-nitrogen removal system. Suitable operating conditions for the hydrotreater include a temperature between about 5 50 and 750 C., a pressure between about 0 and 1000 p.s.i.g., and, preferably, between 50 and 200 p.s.i.g., a liquid hourly space velocity (LHSV) of about 0.1 to 10, and, preferably, 0.5 to 2 and a hydrogen to hydrocarbon feed mol ratio of 1-20/1, and, preferably about 1-5/ 1. The same catalysts used in hydrocracking; namely, an active hydrogenation metal, such as, nickel, platinum, cobalt, molybdenum, tungsten, etc., on an inert oxide support such as alumina, silica, etc., may be used in the hydrotreater. For example, a NiS-MoSz-AlzOa catalyst may be employed at '760 F., 100 p.s.i.g., a LHSV of 0.8, and a mole ratio of H2/ HC of l/l for 9 hours.

Specifically, crude coal tar light oil is fed to the system through line 10 from a source not shown. The feed mixture then passes to distillation column 12 where it is separated to remove a benzene-toluene concentrate or heart-cut boiling between about 140 and 250 F., and, preferably, between 160 and 245 F. This material is discharged through line 14. Any lower boiling material boiling below about 140 F., and, preferably, below 160 F. is discharged through line 16. A material boiling above about 250 F. and preferably above 245 F. is discharged through line 18. Normally the coal tar light oil will contain little or no material boiling below about 140 F. The material from line 14 is then passed to line 97 which is controlled by valve 98. The feed stream of benzenetoluene concentrate passed either through valve 99A and heat exchanger 100A or valve 99B and heat exchanger 100B. From the discharge of heat exchanges 100A or 100B the heat exchanged concentrate then passes through either valve 102A and preheater 103A or valve 102B and preheater 103B, and, thence, to line 32. Line 97 is isolated by Valve 104. Heat exchange in heat exchangers 100A or 100B is effected by passing hot product from line 94 through line 105 valve 106. The hot product from line 105 is then selectively passed through valves 107A and heat exchanger 100A or valve 107B and heat exchanger 100B. From the discharge of heat exchangers 100A or 100B, the heat exchanged product passes to line 42. Line 105 is isolated from line 42 by Valve 108.

The heat exchanged and preheated feed passing through line 97 and thence to line 32 is passed through line 34, and valve 36 to hydrocracking unit 38. Hydrocracking unit 38 is operated under mild hydrocracking conditions, as will be hereinafter pointed out. Hydrogen is supplied to hydrocracking unit 38, from a source not shown, through line 40. Hydrocracked product 38 is discharged through line 42 and passes to line 84. Valve '82, in line 42 is closed and valve 86 in line 84 is opened for this purpose. From line 84, the hydrocracked product from hydrocracker 38 passes to line 32 and thence to line 88 through valve 90 to hydrocracking unit 92. Hydrocracking unit 92 discharges hydrocracked product through line 94 and is supplied with hydrogen from an external source not shown through line 96.

Both hydrocracking units 38 and 92 effect low severity hydrocracking of the material fed to them. Hydrocracking is carried out at temperatures from about 800 to 1200 F., preferably, from about 900 to 1100 F.; at a pressure of about O to 1000 p.s.i.g., and preferably 50 to 200 p.s.i.g.; and at a liquid hourly space velocity (LHSV) of about 0.1 to 10 and preferably 0.5 to 2. Hydrogen is supplied to the hydrocracking units at a hydrogen dihydrocarbon low ratio about 1-5/ 1 and, preferably, 1-2/1. The catalyst in the hydrocracking units may be any active hydrogenation metal such as nickel, platinum, cobalt, molybdenum, tungsten, etc. on an acidic, solid oxide support, such as silica-alumina, boria-alumina, silica-magnesia, uoride promoted alumina, etc.

Normally, severe coking in the heat exchanger section and preheater section of a hydrocracking unit would occur when operating on coal tar light oil and other materials containing heavy ends. However, Aby predistilling the coal tar light oil to produce the benzene-toluene heartcut it has been found that a substantial degree of coking in the heat exchanger and preheater to the hydrocrackers can be eliminated. In addition, the little coking which does occur can be overcome by employing a dual heat exchanger and preheater system shown in the drawing, so that, as plugging from coking occurs in the heat exchanger or the preheater, these can be interchanged for clean out and while still maintaining a continuous operation.

It has been found, in accordance with the present following example, that the particular feed, in this case the benzene-toluene concentrate, can be processed for 24 hours without plugging. The amount of coke laid down in the preheater section was minute and it was indicated that several Weeks processing could be accomplished before plugging occurred. When plugging does occur, by proper manipulation of valves 99, 102 and 107 the heat exchangers and preheaters may be switched at will. The following table shows the results of this operation, utilizaflins and in excess of about fty percent (50%) by volume of benzene and toluene, comprising:

(a) distilling said mixture and removing a benzenetoluene concentrate boiling in the range of about 140 to about 250 F., said concentrate, comprising 0.6% platinum on alumina oxide as a catalyst at a ing benzene, toluene, cyclic diolens, acychc olens, temperature of 940 F., a pressure of 100 p.s.i.g., a liquid cyclo-paraflins and cyclic mono-olens; hourly space velocity (LHSV) of 2 and a hydrogen to (b) hydrogenating the concentrate from (a) ata temhydrocarbon mole ratio of 3 to 1. perature between about 75 and 400 F., a pressure TABLE VI 10 of between about 0 and 3000 p.s.i.g., and in the Fi t s d presence of a rhemum-contaming catalyst;

I'S 80011 Analysis Feed pass pass (C) polymerizing the product obtaioned from (b) at a temperature of about to 300 F., a pressure of gf'mams 3'0 Sjg 7H; about 0 t0 1000 p.s.i.g. and in the presence of a Toluene 1&1 20.4 15 rhenium-containing catalyst Tiu 0.3 0.5 o me (d) distilling the product obtained from (c) and re- TABLE VIII Run No. 1 Run No. 2 Run No. 3 Run No, 4 Polymerlze Hydrocrack Hydrocrack Hydrocrack Catalyst Sto-A120, NiS-WSQ/ Nis-Wsi! Nis-WSQ/ SO2A1203 5102-131203 SiO2Al20s Temperature, F 200 950 950 950 Pressure, p s i Q 100 100 100 200 LHSV 2 0.8 0.8 0.8 Hz/HC, mole/mole 2/1 2/1 2/1 Process time, hrs 30 24 24 4 Product (liquid):

Recovery, percent by volume 96. 5 95 95 95 Benzene, percent by volume 2.5 1. 2 0.6 Benzene, percent by volume.-- 81.5 80. 3 79.6 Toluene, percent by volume- 15. 7 17. 7 18. 9 Toluene, percent by volume 0.3 0.8 0. 9 Thlophene ln benzene, p.p.m- 5-6 0. 75

What is claimed is:

1. A method for recovering benzene and toluene from a benzene-toluene concentrate comprising benzene and toluene in excess of fty percent (50%) by volume, cyclohexadiene and other unsaturated paranic compounds, comprising:

(a) hydrogenating said concentrate at a temperature of about 75 to about 400 F., a pressure of about 0 to about 3000 p.s.i.g. and in the presence of a rhenium-containing catalyst and said catalyst comprising about 0.1 to about 10% rhenium as oxide or sulfide on alumina based on total weight of catalyst;

(b) reforming the product from (a) under a reforming temperature of about 800 to about 1000 P., a pressure of about 0 to about 1000 p.s.i.g., and in the presence of a reforming catalyst comprising platinum or palladium in an amount of 0.1 to 10% by weight on alumina; and

(c) distilling the product from (b) into at least a benzene-toluene fraction.

2. The method of claim 1 wherein prior to the hydrogenating step (a), oxygen compounds present in said concentrate are removed by contacting said concentrate with a molecular sieve.

3. The method of claim 1 wherein the hydrogenation step (a) is conducted at a temperature of about 250 to 300 F.

4. A method for recovering benzene and toluene from normally liquid hydrocarbon mixtures containing cyclic mono-olens, cyclic diolens, acyclic oletins, cyclo-parmoving a concentrate boiling in the range of about to about 250 F.;

(e) hydrocracking the product obtained from (d) at a temperature of 800 to 1200 F., land a pressure of about 0 to about 1000 p.s.i.g. and in the presence of a rhenium-containing catalyst, said catalyst comprising between about 0.1 to 10% by weight rhenium as rhenium oxide or sulde on alumina based on total weight of catalyst; and (f) distilling the product from (e) into at least a benzene-toluene fraction.

References Cited UNITED STATES PATENTS 2,375,464 5/ 1945 Borden 260--674 R 2,953,612 9/ 1960 Haxton et al 260--683.9 2,976,336 3/ 1961 Housam et al. 260-674H 3,294,857 12/ 1966 Tokuhisa et al. 260-674H 2,733,284 1/ 1956 -Hamner 260-674 R 2,849,512 8/ 1958 Banes et al. 260-674R 3,271,297 9/ 1966 Kronig et al 260-674 H 3,296,120 1/ 1967 Doelp et al. 260-674 H 3,400,168 9/ 1968 Fukuda et al. 260-674'H 3,429,804 2/1969 Sze et al 260-674 H 3,449,460 6/ 1969 Tarhan 260-674 A CURTIS R. DAVIS, Primary Examiner U.S. C1. X.R.

260-674 A, 674 H, 683.9, 683.15 R; 208-57, 60, 61

` -UNTTED STATES PATENT OFFICE CERTIFICATE OF CORRECTION Patent No. 3,806,553 Dated April 23, 1974 Invent-ONS) Stephen M. Kovach and Ralph E. Patrick It is certified that error appears in the above-identified patent and vthat said Letters Patent are hereby corrected as shown below:`

Column 1, line 29, "acatalyst" should read --a catalyst". Column 2, line 31, "quit" should read "quite". y E Column 4, line 43,4 delete "and, preferably between about 0.5 to 2. Hydrogen, at a" and substitute therefor. --,and

preferably, about 100 to 300 cubic feet per barrel, is. Column 4, VTable I, under Run No. 2 and opposite Pressure, p.s.i.g. insert a zero -0.

Column 5, line 24, "techninue" should read -technique-. Column 6, line 44, "Z50-F" should read -2500F. Column 6, line 47, delete "have a boiling range of about 140 to 25001:."

` l and substitute therefor --l6 and feed mixture boiling Collmln 9, line 60, n "Thhe'v' should read The f Column 9, line 64, "7500C" should read 7500F. Column *11, delete Table VIII.

" Signed en d sealed this 7th day of January 19.75.

(SEAL) Attest:

MecoY M. GIBSON JR. Y c. MARSHALL DANN Attestng Officer Commissioner off Patents F ORM PO-1050 (\0-69) -USCOMM-DC 60ans-P69 u.s. GovEnNMzNr PRINTING africa; n o-sn-su 

